Method for the dehydrogenation of hydrocarbons

ABSTRACT

In a process for the heterogeneously catalyzed dehydrogenation in one or more reaction zones of one or more dehydrogenatable C 2 -C 30 -hydrocarbons in a reaction gas mixture comprising them, with at least part of the heat of dehydrogenation required being generated directly in the reaction gas mixture in at least one reaction zone by combustion of hydrogen, the hydrocarbon or hydrocarbons and/or carbon in the presence of an oxygen-containing gas, the reaction gas mixture comprising the dehydrogenatable hydrocarbon or hydrocarbons is brought into contact with a Lewis-acid dehydrogenation catalyst which has essentially no Brönsted acidity.

The present invention relates to a process for the heterogeneouslycatalyzed dehydrogenation of dehydrogenatable C₂-C₃₀-hydrocarbons.

Dehydrogenated hydrocarbons are required in large quantities as startingmaterials for numerous industrial processes. For example, dehydrogenatedhydrocarbons are used in the production of detergents, antiknockgasoline and pharmaceutical products. Likewise, numerous plastics areproduced by polymerization of olefins.

For example, acrylonitrile, acrylic acid or C₄ oxo alcohols are preparedfrom propylene. Propylene is at present produced predominantly by steamcracking or catalytic cracking of suitable hydrocarbons or hydrocarbonmixtures such as naphtha.

Propylene can also be prepared by heterogeneously catalyzeddehydrogenation of propane.

To obtain acceptable conversions in heterogeneously catalyzeddehydrogenations even on a single pass through the reactor, relativelyhigh reaction temperatures generally have to be employed. Typicalreaction temperatures for gas-phase dehydrogenations are from 300 to700° C. In general, one molecule of hydrogen is produced per molecule ofhydrocarbon.

The dehydrogenation of hydrocarbons proceeds endothermically. The heatof dehydrogenation necessary for obtaining a desired conversion has tobe introduced into the reaction gas either beforehand and/or during thecourse of the catalytic dehydrogenation. In most known dehydrogenationprocesses, the heat of dehydrogenation is generated outside the reactorand is introduced into the reaction gas from the outside. However, thisrequires complicated reactor and process concepts and leads,particularly at high conversions, to steep temperature gradients in thereactor, which is accompanied by the risk of increased byproductformation. Thus, for example, a plurality of adiabatic catalyst beds canbe arranged in annular gap reactors connected in series. The reactiongas mixture is superheated by means of heat exchangers on its way fromone catalyst bed to the next catalyst bed and cools down again duringpassage through the subsequent reactor. To achieve high conversionsusing such a reactor concept, it is necessary either to increase thenumber of reactors connected in series or to increase the reactor inlettemperature of the gas mixture. The resulting superheating inevitablyleads to increased byproduct formation due to cracking reactions.Another known method is to arrange the catalyst bed in a tube reactorand to generate the heat of dehydrogenation by burning combustible gasesoutside the tube reactor and introduce it into the interior of thereactor via the tube wall. In these reactors, high conversions lead tosteep temperature gradients between the wall and the interior of thereation tube.

An alternative is to generate the heat of dehydrogenation directly inthe reaction gas mixture of the dehydrogenation by oxidation of hydrogenformed in the dehydrogenation or additionally fed in or of hydrocarbonspresent in the reaction gas mixture by means of oxygen. For thispurpose, an oxygen-containing gas and possibly hydrogen is/are added tothe reaction gas mixture either upstream of the first catalyst bed orupstream of subsequent catalyst beds. The heat of reaction liberated inthe oxidation also prevents high temperature gradients in the reactor athigh conversions. At the same time, a very simple process concept isrealized by omission of the indirect heating of the reactor.

U.S. Pat. No. 4,788,371 describes a process for steam dehydrogenation ofdehydrogenatable hydrocarbons in the gas phase combined with oxidativereheating of the intermediates, with the same catalyst being used forthe selective oxidation of hydrogen and the steam dehydrogenation. Here,hydrogen can be introduced as co-feed. The catalyst used comprises anoble metal of group VIII, an alkali metal and a further metal selectedfrom the group consisting of B, Ga, In, Ge, Sn and Pb on an inorganicoxide support such as aluminum oxide. The process can be carried out inone or more stages in a fixed or moving bed.

WO 94/29021 describes a catalyst which comprises a support consistingessentially of a mixed oxide of magnesium and aluminum Mg(Al)O and alsoa noble metal of group VIII, preferably platinum, a metal of group IVA,preferably tin, and possibly an alkali metal, preferably cesium. Thecatalyst is used in the dehydrogenation of hydrocarbons, which can becarried out in the presence of oxygen.

U.S. Pat. No. 5,733,518 describes a process for the selective oxidationof hydrogen by oxygen in the presence of hydrocarbons such as n-butaneover a catalyst comprising a phosphate of germanium, tin, lead, arsenic,antimony or bismuth, preferably tin. The combustion of the hydrogengenerates, in at least one reaction zone, the heat of reaction necessaryfor the endothermic dehydrogenation.

EP-A 0 838 534 describes a catalyst for the steam-free hydrogenation ofalkanes, in particular isobutane, in the presence of oxygen. Thecatalyst used comprises a platinum group metal applied to a supportcomprising tin oxide/zirconium oxide and having a tin content of atleast 10%. The oxygen content of the feed stream for the dehydrogenationis calculated so that the quantity of heat generated by the combustionreaction of hydrogen and oxygen is equal to the quantity of heatrequired for the dehydrogenation.

WO 96/33151 describes a process for the dehydrogenation of aC₂-C₅-alkane in the absence of oxygen over a dehydrogenation catalystcomprising Cr, Mo, Ga, Zn or a group VIII metal with simultaneousoxidation of the resulting hydrogen over a reducible metal oxide, e.g.an oxide of Bi, In, Sb, Zn, Tl, Pb or Te. The dehydrogenation has to beinterrupted at regular intervals in order to reoxidize the reduced oxideby means of an oxygen source. U.S. Pat. No. 5,430,209 describes acorresponding process in which the dehydrogenation step and theoxidation step proceed sequentially and the associated catalysts areseparated physically from one another. Catalsyts used for the selectiveoxidation of hydrogen are oxides of Bi, Sb and Te and also their mixedoxides.

Finally, WO 96/33150 describes a process in which a C₂-C₅-alkane isdehydrogenated over a dehydrogenation catalyst in a first stage, theoutput gas from the dehydrogenation stage is mixed with oxygen and, in asecond stage, passed over an oxidation catalyst, preferably Bi₂O₃, so asto selectively oxidize the hydrogen formed to water, and, in a thirdstage, the output gas from the second stage is again passed over adehydrogenation catalyst.

The catalyst system used has to meet demanding requirements in respectof achievable alkane conversion, selectivity to formation of alkenes,mechanical stability, thermal stability, carbonization behavior,deactivation behavior, regenerability, stability in the presence ofoxygen and insensitivity to catalyst poisons such as CO, sulfur- andchlorine-containing compounds, alkynes, etc., and economics.

The catalysts of the prior art do not meet these requirements,particularly in respect of the achievable conversions and selectivities,operating lives and regenerability, to a satisfactory extent.

It is an object of the present invention to provide a process for thedehydrogenation of hydrocarbons which ensures high conversions,space-time yields and selectivities.

We have found that this object is achieved by a process for theheterogeneously catalyzed dehydrogenation in one or more reaction zonesof one or more dehydrogenatable C₂-C₃₀-hydrocarbons in a reaction gasmixture comprising them, with at least part of the heat ofdehydrogenation required being generated directly in the reaction gasmixture in at least one reaction zone by combustion of hydrogen, thehydrocarbon or hydrocarbons and/or carbon in the presence of anoxygen-containing gas, wherein the reaction gas mixture comprising thedehydrogenatable hydrocarbon or hydrocarbons is brought into contactwith a Lewis-acid dehydrogenation catalyst which has essentially noBrönsted acidity.

The dehydrogenation catalyst used according to the present invention hasessentially no Brönsted acidity, but a high Lewis acidity. Thedetermination of the Lewis and Brönsted acidities of the dehydrogenationcatalysts is carried out by adsorption of pyridine as basic probemolecule on the activated catalyst with subsequent quantitativeFT-IR-spectrometric determination of the Brönsted- and Lewis-specificadsorbates. This method makes use of the fact that the adsorbed probemolecules give different IR spectra depending on whether they are boundto a Brönsted center or a Lewis center. At the Brönsted center, protontransfer takes place to form a local ion pair with the pyridinium ion ascation. The adsorbed pyridinium ion displays a Brönsted-specificabsorption band at 1545 cm⁻¹ in the IR spectrum. At the Lewis center, onthe other hand, the probe molecule pyridine is coordinated via its freeelectron pair on the ring nitrogen to the electron-deficient center.This results in an IR spectrum different to that of the Brönstedadsorbate. The Lewis band is found at 1440 cm⁻¹. Quantitative evaluationof the Brönsted and Lewis bands enables the Brönsted and Lewis centersto be determined separately. The band assignment is based on the work ofTurkevich (C. H. Kline, J. Turkevich: J. Chem. Phys. 12, 300 (1994)).

The assignment of the resulting pyridine bands in the FT-IR spectrum isas follows:

-   Lewis (L): 1440 cm⁻¹-   Brönsted (B): 1545 cm⁻¹-   Control band B+L: 1490 cm⁻¹-   Physisorbed pyridine: 1590 cm⁻¹ (additionally 1440 cm⁻¹)

The transmission cell used for the measurements is a reconstruction ofthe prototype of Gallei and Schadow (E. Gallei et al.: Rev. Sci.Instrur. 45 (12), 1504 (1976)). The cell comprises a stainless steelbody with parallel IR-transparent windows made of CaF₂. The intrinsicabsorption of the windows makes it possible to measure only in aspectral range of about 1200-4000 cm⁻¹. A circuit for cooling or heatingfluid is provided in the cell body. In the lid of the cell there is asolid, plate-shaped sample holder with built-in cartridge heating (400°C.). The self-supporting sample compound is laid in an annular doubletemplate and screwed into the heating plate, and the cell lid is screwedonto the cell body. The measurement cell can be evacuated to 10⁻⁵-10⁻⁶mbar.

For sample preparation, the catalyst material is ground finely in amortar and pressed between two stainless steel plates with micaunderlays in a film press at a pressing pressure of 50 kN to give aself-supporting wafer. The thickness depends on the intrinsic IRabsorption of the material and is typically in the range from 30 to 100μm. Pellets having a diameter of about 5 mm are cut from the wafer.

The activation of the sample in the measurement cell is carried out inair at 390° C. After heating, the cell is evacuated to 10⁻⁵-10⁻⁶ mbar.It is then cooled to the gas treatment temperature of 80° C. under ahigh vacuum.

The sample is subsequently treated with gaseous pyridine at a pressurewhich can be from 10⁻² to 3 mbar. Control spectra of the sample withadsorbate are recorded until a steady adsorption state has beenestablished at the gas treatment pressure concerned. The cell issubsequently evacuated to a high vacuum (10⁻⁵ mbar). This removesphysisorbates. After evacuation is complete, adsorbate spectra arerecorded.

To determine the Lewis and Brönsted acidities, the intensities of thebands at 1440 cm⁻¹ and 1545 cm⁻¹ obtained at a particular thickness ofthe sample and a set equilibrium pressure of pyridine are evaluated incomparison with one another. If no band at 1545 cm⁻¹ is discernible (noBrönsted acidity), the band at 1490 cm⁻¹ can also be employed fordetermining the Lewis acidity.

The measured absorbances are expressed as a ratio to the thickness ofthe sample (in integrated extinction units (IEE) per μm of thickness).The single-beam spectrum of the sample which has not been treated withpyridine gas (cooled to 80° C.) under high vacuum serves as backgroundof the adsorbate spectra. This completely balances out matrix bands.

1 AU corresponds to one thousand times the measured absorbance (reportedin integrated extinction units IEE) divided by the thickness of thesample (in μm) obtained in the determination of the Lewis and Brönstedacidities of the dehydrogenation catalysts using the probe gas pyridine.

The dehydrogenation catalysts used according to the present inventiongenerally have no detectable Brönsted acidity, i.e. their Brönstedacidity is less than 0.1 AU. However, they have a high Lewis acidity.The Lewis acidity of the dehydrogenation catalysts is generally greaterthan 1 AU, preferably greater than 3 AU, particularly preferably greaterthan 6 AU.

The dehydrogenation catalysts used according to the present inventiongenerally comprise a support and an active composition. The supportcomprises a heat-resistant oxide or mixed oxide. The dehydrogenationcatalyst preferably comprises a metal oxide selected from the groupconsisting of zirconium dioxide, zinc oxide, aluminum oxide, silicondioxide, titanium dioxide, magnesium oxide, lanthanum oxide, ceriumoxide and mixtures thereof as support. Preferred supports are zirconiumdioxide and/or silicon dioxide; particular preference is given tomixtures of zirconium dioxide and silicon dioxide.

The active composition of the dehydrogenation catalyst used according tothe present invention generally comprises one or more elements oftransition group VIII, preferably platinum and/or palladium,particularly preferably platinum. In addition, the dehydrogenationcatalyst can further comprise one or more elements of main groups Iand/or II, preferably potassium and/or cesium. The dehydrogenationcatalyst may also further comprise one or more elements of transitiongroup III including the lanthanides and actinides, preferably lanthanumand/or cerium. Finally, the dehydrogenation catalyst can furthercomprise one or more elements of main groups III and/or IV, preferablyone or more elements selected from the group consisting of boron,gallium, silicon, germanium, tin and lead, particularly preferably tin.

In a preferred embodiment, the dehydrogenation catalyst comprises atleast one element of transition group VIII, at least one element of maingroups I and/or II, at least one element of main groups III and/or IVand at least one element of transition group III including thelanthanides and actinides.

To produce the dehydrogenation catalysts used according to the presentinvention, it is possible to use precursors of oxides of zirconium,silicon, aluminum, titanium, magnesium, lanthanum or cerium which can beconverted into the oxides by calcination. These can be produced by knownmethods, for example by the sol-gel process, precipitation of salts,dehydration of the corresponding acids, dry mixing, slurrying or spraydrying. For example, to produce a ZrO₂.Al₂O₃.SiO₂ mixed oxide, awater-rich zirconium oxide of the formula ZrO₂.xH₂O can firstly beprepared by precipitation of a suitable zirconium-containing precursor.Suitable precursors of zirconium are, for example, Zr(NO₃)₄, ZrOCl₂ orZrC₄. The precipitation itself is carried out by addition of a base suchas NaOH, KOH, Na₂CO₃ or NH₃ and is described, for example, in EP-A 0 849224.

To produce a ZrO₂.SiO₂ mixed oxide, the zirconium-containing precursorobtained above can be mixed with a silicon-containing precursor.Well-suited precursors of SiO₂ are, for example, water-containing solsof SiO₂ such as Ludox™. The two components can be mixed, for example, bysimple mechanical mixing or by spray drying in a spray dryer.

To produce a ZrO₂.SiO₂.Al₂O₃ mixed oxide, the SiO₂.ZrO₂ powder mixtureobtained as described above can be admixed with an aluminum-containingprecursor. This can be achieved, for example, by simple mechanicalmixing in a kneader. However, the ZrO₂.SiO₂.Al₂O₃ mixed oxide can alsobe produced in a single step by dry mixing the individual precursors.

The supports for the dehydrogenation catalysts used according to thepresent invention have, inter alia, the advantage that they can easilybe shaped. For this purpose, the powder mixture obtained is admixed witha concentrated acid in a kneader and then converted into a shaped body,e.g. by means of a ram extruder or a screw extruder.

The dehydrogenation catalysts used according to the present inventionhave, in particular embodiments, a defined pore structure. When usingmixed oxides, it is possible to influence the pore structure in atargeted manner. The particle sizes of the various precursors influencethe pore structure. Thus, for example, macropores can be generated inthe microstructure by use of Al₂O₃ having a low loss on ignition and adefined particle size distribution. In this context, the use of Al₂O₃having a loss on ignition of about 3% (e.g. Puralox®) has been found tobe useful.

A further possible way of producing supports having specific pore radiusdistributions for the dehydrogenation catalysts used according to thepresent invention is to add various polymers during production of thesupport and subsequently to remove them completely or partly bycalcination so as to form pores in defined pore ranges. Mixing thepolymers and the oxide precursors can be carried out, for example, bysimple mechanical mixing or by spray drying in a spray dryer.

The use of PVP (polyvinylpyrrolidone) has been found to be particularlyuseful for producing supports having a bimodal pore radius distribution.If PVP is added to one or more oxide precursors of oxides of theelements Zr, Ti, Al or Si in a production step, macropores having sizesin the range from 200 to 5000 nm are formed after calcination. A furtheradvantage of the use of PVP is that it makes the support easier toshape. Thus, extrudates having good mechanical properties can readily beproduced from freshly precipitated hydrous ZrO₂.xH₂O which haspreviously been dried at 120° C. when PVP and formic acid are added,even without further oxide precursors.

The calcination of the supports for the dehydrogenation catalysts usedaccording to the present invention is advantageously carried out afterapplication of the active components and is carried out at from 400 to1000° C., preferably from 500 to 700° C., particularly preferably from550 to 650° C. and in particular at from 560 to 620° C. The calcinationtemperature should usually be at least as high as the reactiontemperature of the dehydrogenation in which the dehydrogenationcatalysts are used according to the present invention.

The supports of the dehydrogenation catalysts used according to thepresent invention generally have high BET surface areas aftercalcination. The BET surface areas are generally greater than 40 m²/g,preferably greater than 50 m²/g, particularly preferably greater than 70m²/g. The pore volume of the dehydrogenation catalysts used according tothe present invention is usually from 0.2 to 0.6 ml/g, preferably from0.25 to 0.5 ml/g. The mean pore diameter of the dehydrogenationcatalysts used according to the present invention, which can bedetermined by Hg porosimetry, is from 3 to 20 nm, preferably from 4 to15 nm.

Furthermore, the dehydrogenation catalysts used according to the presentinvention have a bimodal pore radius distribution. The pores have sizesin the range up to 20 nm and in the range from 40 to 5000 nm. Thesepores all together make up at least 70% of the total pore volume of thedehydrogenation catalyst. The proportion of pores smaller than 20 nm isgenerally in the range from 20 to 60%, while the proportion of pores inthe range from 40 to 5000 nm is generally likewise from 20 to 60%.

The dehydrogenation-active component, which is usually a metal oftransition group VIII, is generally applied by impregnation with asuitable metal salt precursor. Instead of impregnation, thedehydrogenation-active component can also be applied by other methods,for example spraying the metal salt precursor onto the support. Suitablemetal salt precursors are, for example, the nitrates, acetates andchlorides of the corresponding metals; complex anions of the metals usedare also possible. Preference is given to using platinum as H₂PtCl₆ orPt(NO₃)₂. Suitable solvents for the metal salt precursors include bothwater and organic solvents. Particularly useful solvents are water andlower alcohols such as methanol and ethanol.

When using noble metals as dehydrogenation-active components, suitableprecursors also include the corresponding noble metal sols which can beprepared by one of the known methods, for example by reduction of ametal salt using a reducing agent in the presence of a stabilizer suchas PVP. The method of preparation is dealt with comprehensively in theGerman Patent Application DE 195 00 366.

The amount of a noble metal present as dehydrogenation-active componentin the dehydrogenation catalysts used according to the present inventionis from 0 to 5% by weight, preferably from 0.05 to 1% by weight,particularly preferably from 0.05 to 0.5% by weight.

The further components of the active composition can be applied eitherduring production of the support, for example by coprecipitation, orsubsequently, for example by impregnating the support with suitableprecursor compounds. Precursor compounds used are generally compoundswhich can be converted into the corresponding oxides by calcination.Suitable precursors are, for example, hydroxides, carbonates, oxalates,acetates, chlorides or mixed hydroxycarbonates of the correspondingmetals.

In advantageous embodiments, the active composition further comprisesthe following additional components:

-   -   at least one element of main group I or II, preferably cesium        and/or potassium, in an amount of from 0 to 20% by weight,        preferably from 0.1 to 15% by weight, particularly preferably        from 0.2 to 10% by weight;    -   at least one element of transition group III including the        lanthanides and actinides, preferably lanthanum and/or cerium,        in an amount of from 0 to 20% by weight, preferably from 0.1 to        15% by weight, particularly preferably from 0.2 to 10% by        weight;    -   at least one element of main groups III and IV, preferably tin,        in an amount of from 0 to 10% by weight.

The dehydrogenation catalyst is preferably halogen-free.

The dehydrogenation catalyst can be used in the form of a fixed bed inthe reactor or, for example, in the form of a fluidized bed and can havean appropriate shape. Suitable shapes are, for example, granules,pellets, monoliths, spheres or extrudates (rods, wagon wheels, stars,rings).

As dehydrogenatable hydrocarbons, it is possible to use paraffins,alkylaromatics, naphthenes or olefins having from 2 to 30 carbon atoms.The process is particularly useful for the dehydrogenation ofstraight-chain or branched hydrocarbons having a chain length of from 2to 15 carbon atoms, preferably from 2 to 5 carbon atoms. Examples areethane, propane, n-butane, isobutane, n-pentane, isopentane, n-hexane,n-heptane, n-octane, n-nonane, n-decane, n-undecane, ndodecane,n-tridecane, n-tetradecane and n-pentadecane. The particularly preferredhydrocarbon is propane. In the further description of the invention, thediscussion will frequently concern this particularly preferred case ofpropane dehydrogenation, but the corresponding features applyanalogously to other dehydrogenatable hydrocarbons.

Since the dehydrogenation reaction is accompanied by an increase involume, the conversion can be increased by lowering the partial pressureof the reactants. This can be achived in a simple manner by, forexample, dehydrogenation under reduced pressure and/or by mixing in aninert gas. Suitable inert gases are, for example, nitrogen, steam,carbon dioxide and noble gases such as He, Ne or Ar. Preference is givento diluents which are inert under the reaction conditions (i.e. diluentswhich are changed chemically to an extent of less than 5 mol %,preferably less than 3 mol % and even better less than 1 mol %). Afurther advantage of dilution with steam is generally reducedcarbonization of the dehydrogenation catalyst used according to thepresent invention and thus an increased operating life, since the steamreacts with carbon formed according to the principle of carbongasification. The ratio of steam to the hydrocarbon to be dehydrogenatedis in the range from 0 to 10 mol/mol, preferably from 0.1 to 5 mol/mol.

The process of the present invention is carried out in at least onereaction zone with simultaneous generation of heat by exothermicreaction of hydrogen, hydrocarbon and/or carbon in the presence of anoxygen-containing gas. In general, the total amount of oxygenintroduced, based on the total amount of the hydrocarbon to bedehydrogenated, is from 0.001 to 0.5 mol/mol, preferably from 0.005 to0.2 mol/mol, particularly preferably from 0.05 to 0.2 mol/mol. Ingeneral, the amount of oxygen-containing gas added to the reaction gasmixture is chosen so that the combustion of the hydrogen or hydrocarbonpresent in the reaction gas mixture and/or the carbon present in theform of carbon deposits generates the quantity of heat required fordehydrogenation of the hydrocarbon to the alkene. In particularembodiments, the quantity of heat generated by the combustion reactionwith oxygen can also be greater or lesser than the quantity of heatrequired for the dehydrogenation of the hydrocarbon. Oxygen can be usedeither as pure oxygen or in admixture with inert gases such as CO₂, N₂or noble gases. Air is particularly preferred as oxygen-containing gas.As an alternative to molecular oxygen, it is also possible to usefurther oxygen-containing gaseous oxidants, for example dinitrogen oxideor ozone. The inert gases and the resulting combustion gases generallyhave an additional diluting effect and thus promote the heterogeneouslycatalyzed dehydrogenation.

The hydrogen burnt for heat generation can be the hydrogen formed in thedehydrogenation or be additional hydrogen added to the reaction gasmixture.

In one embodiment of the invention, no hydrogen is added to the reactiongas mixture and the heat required for dehydrogenation is generated atleast partly by combustion (exothermic reaction) of hydrocarbon and ofthe hydrogen formed in the dehydrogenation.

In a further embodiment, additional hydrogen is added to the reactiongas mixture.

The dehydrogenation catalyst used according to the present inventiongenerally also catalyzes the combustion of hydrocarbons and of hydrogenwith oxygen, so that no specific oxidation catalyst different from thisis necessary in principle. In one embodiment, a specific, differentoxidation catalyst which selectively carbonizes the oxidation ofhydrogen for the generation of heat is used in addition to thedehydrogenation catalyst, particularly when additional hydrogen isadded.

If, as in one embodiment of the invention, no additional hydrogen isadded to the reaction gas mixture, the heat of dehydrogenation canreadily be generated by catalytic combustion of the hydrocarbons and ofhydrogen formed in the dehydrogenation over the dehydrogenationcatalyst. Suitable, oxygen-insensitive dehydrogenation catalysts whichcatalyze the combustion of hydrocarbons are the above-describedLewis-acid catalysts. Preference is given to dehydrogenation catalystsof the type described above which comprise at least one element oftransition group VIII, at least one element of main groups I and/or II,at least one element of main groups III and/or IV and at least oneelement of transition group III including the lanthanides and actinideson zirconium oxide and/or silicon dioxide as support.

In a preferred embodiment, hydrogen is added to the reaction gas mixturefor direct heat generation by combustion. In general, the amount ofhydrogen added to the reaction gas mixture is such that the molar ratioof H₂/O₂ in the reaction gas mixture immediately after the addition isfrom 0.1 to 200 mol/mol, preferably from 1 to 20 mol/mol, particularlypreferably from 2 to 10 mol/mol. In the case of multistage reactors,this applies to each intermediate introduction of hydrogen and oxygen.

The combustion of hydrogen occurs catalytically. In one embodiment ofthe invention, no specific oxidation catalyst different from thedehydrogenation catalyst is used. In a particularly preferredembodiment, the reaction is carried out in the presence of one or moreoxidation catalysts which selectively catalyze the combustion reactionof hydrogen and oxygen in the presence of hydrocarbons. As a result, thecombustion reaction of hydrocarbons with oxygen to give CO and CO₂proceeds only to a subordinate degree, which has a significant positiveeffect on the achieved selectivities for the formation of alkenes. Thedehydrogenation catalyst and the oxidation catalyst are preferablypresent in different reaction zones.

In a multistage reaction, the oxidation catalyst can be present in onlyone reaction zone, in a plurality of reaction zones or in all reactionzones.

The catalyst which selectively catalyzes the oxidation of hydrogen inthe presence of hydrocarbons is preferably located at points at whichhigher oxygen partial pressures prevail than at other points of thereactor, in particular in the vicinity of the feed point for theoxygen-containing gas. Oxygen-containing gas and/or hydrogen can be fedin at one or more points on the reactor.

A preferred catalyst for the selective combustion of hydrogen comprisesoxides or phosphates selected from the group consisting of the oxidesand phosphates of germanium, tin, lead, arsenic, antimony and bismuth.

A further, preferred catalyst for the catalytic combustion of hydrogencomprises a noble metal of transition group VIII or I.

In heterogeneously catalyzed dehydrogenations of hydrocarbons, smallamounts of high-boiling, high molecular weight organic compounds orcarbon are generally formed over time, and these deposit on the catalystsurface and deactivate the catalyst as time goes on. The dehydrogenationcatalysts used according to the present invention have a low tendency tosuffer from carbonization and have a low deactivation rate.

The dehydrogenation catalysts used according to the present inventionmake it possible to achieve high space-time yields which for thedehydrogenation of propane are above 2 kg of propene/kg of catalyst*hand are thus significantly above the space-time yields of the processesof the prior art. Diluting the reaction gas mixture with inert gas,increasing the reaction temperature and/or lowering the reactionpressure enable the thermodynamically possible limiting conversions tobe increased so far that they are significantly above the reactionconversions sought. In this way, space-time yields of over 6 kg ofpropene/kg of catalyst*h can be achieved in the presence of thecatalysts used according to the present invention.

The space velocity (GHSV) over the catalyst in this operating mode alsoreferred to as high-load operation can be >8000 h⁻¹.

Regeneration of the dehydrogenation catalyst can be carried out usingmethods known per se. Thus, as described above, steam can be added tothe reaction gas mixture. The deposited carbon is partly or completelyremoved under these reaction conditions according to the principle ofcarbon gasification.

As an alternative, an oxygen-containing gas can be passed at hightemperature over the catalyst bed from time to time so as to burn offthe deposited carbon.

After a prolonged period of operation, the dehydrogenation catalyst usedaccording to the present invention is preferably regenerated by, at atemperature of from 300 to 600° C., frequently at from 350 to 500° C.,firstly carrying out a flushing operation with inert gas andsubsequently, in a first regeneration step, passing air diluted withnitrogen over the catalyst bed. The space velocity over the catalyst ispreferably from 50 to 10 000 h⁻¹ and the oxygen content is from about0.5 to 2% by volume. In subsequent regeneration steps, the oxygencontent is gradually increased to about 20% by volume (pure air).Preference is given to carrying out from 2 to 10, particularlypreferably from 2 to 5, regeneration steps. In general, the catalyst issubsequently regenerated further using pure hydrogen or hydrogen dilutedwith an inert gas (hydrogen content >1% by volume) under otherwiseidentical conditions. All regeneration steps are preferably carried outin the presence of water vapor.

The process of the present invention can in principle be carried out inall reactor types known from the prior art and by all operatingprocedures known from the prior art. The additional introduction ofoxygen leads to at least part of the heat of reaction or of the energyrequired for heating the reaction gas mixture being supplied by directcombustion and not having to be introduced indirectly via heatexchangers.

A comprehensive description of suitable reactor types and operatingprocedures is also given in “Catalytica® Studies Division, OxidativeDehydrogenation and Alternative Dehydrogenation Processes, Study Number4192 OD, 1993, 430 Ferguson Drive, Mountain View, Calif., 94043-5272U.S.A.”.

A suitable form of reactor is a fixed-bed tube or multitube(shell-and-tube) reactor. In this, the catalyst (dehydrogenationcatalyst and, if desired, specific oxidation catalyst) is present as afixed bed in a reaction tube or in a bundle of reaction tubes. Thereaction tubes are customarily heated indirectly by a gas, e.g. ahydrocarbon such as methane, being burnt in the space surrounding thereaction tubes. In such a reactor, it is advantageous to employ thisindirect form of heating only for the first about 20-30% of the lengthof the fixed bed and to heat the remaining length of the bed to therequired reaction temperature by means of the radiative heat liberatedby the indirect heating. According to the present invention, theindirect heating of the reaction gas can advantageously be coupled withthe direct heating by combustion in the reaction gas mixture. Couplingthe direct introduction of heat with the indirect introduction of heatmakes it possible to achieve approximately isothermal reactionconditions. Customary internal diameters of the reaction tubes are fromabout 10 to 15 cm. A typical shell-and-tube reactor used fordehydrogenation has from about 300 to 1000 reaction tubes. Thetemperature in the interior of the reaction tube is generally in therange from 300 to 700° C., preferably from 400 to 700° C. The operatingpressure is usually from 0.5 to 8 bar, frequently from 1 to 2 bar whenusing low steam dilution (corresponding to the BASF Linde process) butalso from 3 to 8 bar when using a high steam dilution (corresponding tothe steam active reforming process (STAR process) of Phillips PetroleumCo., cf. U.S. Pat. No. 4,902,849, U.S. Pat. No. 4,996,387 and U.S. Pat.No. 5,389,342). In general, the product mixture leaves the reaction tubeat a from 50 to 100° C. lower temperature. Typical space velocities overthe catalyst in the case of propane are from 500 to 2000 h⁻¹. Thecatalyst geometry can be, for example, spherical or cylindrical (hollowor solid).

The process of the present invention can also be carried out in a movingbed reactor. For example, the moving catalyst bed can be accommodated ina radial flow reactor. In this, the catalyst slowly moves from the topdownward, while the reaction gas mixture flows radially. This method ofoperation is employed, for example, in the UOP-Oleflex dehydrogenationprocess. Since the reactors are operated pseudoadiabatically in thisprocess, it is advantageous to employ a plurality of reactors connectedin series (typically up to four reactors). Before or in each reactor,the inflowing gas mixture is heated to the required reaction temperatureby combustion in the presence of the oxygen fed in. The use of aplurality of reactors makes it possible to avoid large differences inthe temperatures of the reaction gas mixture between reactor inlet andreactor outlet while nevertheless achieving high total conversions.

When the catalyst has left the moving bed reactor, it is passed toregeneration and subsequently reused. The dehydrogenation catalyst usedaccording to the present invention generally has a spherical shape.Hydrogen can also be added to the hydrocarbon to be dehydrogenated,preferably propane, to avoid rapid catalyst deactivation. The operatingpressure is typically from 2 to 5 bar. The molar ratio of hydrogen topropane is preferably from 0.1 to 10. The reaction temperatures arepreferably from 550 to 660° C.

A hydrocarbon dehydrogenation by the process of the present inventioncan also, as described in Chem. Eng. Sci. 1992 b, 47 (9-11) 2313, becarried out in the presence of a heterogeneous catalyst in a fluidizedbed, with the hydrocarbon not being diluted. In this method, twofluidized beds are advantageously operated in parallel, with one of themgenerally being in the regeneration mode. The operating pressure istypically from 1 to 2 bar, and the dehydrogenation temperature isgenerally from 550 to 600° C. The heat necessary for the dehydrogenationis introduced into the reaction system by preheating the dehydrogenationcatalyst to the reaction temperature. The use according to the presentinvention of an oxygen-containing co-feed makes it possible to omit thepreheater and to generate the necessary heat directly in the reactorsystem by combustion in the presence of oxygen.

In a particularly preferred embodiment of the process of the presentinvention, the dehydrogenation is carried out in a tray reactor. Thiscontains one or more successive catalyst beds. The number of catalystbeds can be from 1 to 20, advantageously from 2 to 8, in particular from4 to 6. The reaction gas preferably flows radially or axially throughthe catalyst beds. In general, such a tray reactor is operated using afixed catalyst bed.

In the simplest case, the fixed catalyst beds are arranged axially or inthe annular gaps of concentric, upright mesh cylinders in a shaftfurnace reactor. One shaft furnace reactor corresponds to one tray. Itis possible for the process of the present invention to be carried outin a single shaft furnace reactor, but this is less preferred.

In an operating mode without oxygen as co-feed, the reaction gas mixtureis subjected to intermediate heating on its way from one catalyst bed tothe next catalyst bed in the tray reactor, e.g. by passing it over heatexchanger ribs heated by means of hot gases or by passing it throughtubes heated by means of hot combustion gases.

In the process of the present invention, the above-describedintermediate heating is carried out at least partly by direct means. Forthis purpose, a limited amount of molecular oxygen is added to thereaction gas mixture either before it flows through the first catalystbed and/or between the subsequent catalyst beds. Thus, hydrocarbonspresent in the reaction gas mixture, carbon or carbon-like compoundswhich have deposited on the catalyst surface and also hydrogen formedduring the dehydrogenation are burnt to a limited extent over thecatalyst used according to the present invention. The heat of reactionliberated in this combustion thus makes it possible for theheterogeneously catalyzed hydrocarbon dehydrogenation to be operatedvirtually isothermally. The process can be operated with or withoutintroduction of additional hydrogen.

In one embodiment of the invention, intermediate introduction ofoxygen-containing gas and possibly hydrogen is carried out upstream ofeach tray of the tray reactor. In a further embodiment of the process ofthe present invention, the introduction of oxygen-containing gas andpossibly hydrogen is carried out upstream of each tray apart from thefirst tray. In a preferred embodiment, intermediate introduction ofhydrogen is employed; in a specific embodiment of this, a bed of aspecific oxidation catalyst is present downstream of each introductionpoint and is followed by a bed of the dehydrogenation catalyst, and in asecond specific embodiment, no specific oxidation catalyst is present.In a further preferred embodiment, no hydrogen is introduced.

The dehydrogenation temperature is generally from 400 to 800° C. and thepressure is generally from 0.2 to 5 bar, preferably from 0.5 to 2 bar,particularly preferably from 1 to 1.5 bar. The space velocity (GHSV) isgenerally from 500 to 2000 h⁻¹ and in high-load operation up to 16000h⁻¹, preferably from 4000 to 16000 h⁻¹.

The hydrocarbon used in the process of the present invention does nothave to be a pure compound. Rather, the hydrocarbon used can compriseother dehydrogenatable gases such as methane, ethane, ethylene, propane,propene, butanes, butenes, propyne, acetylene, H₂S or pentanes. Inparticular, the dehydrogenation of the present invention can also becarried out using alkane mixtures which are produced industrially andare available in large quantities, for example LPG (liquefied petroleumgas). It is also possible to use circulation gases originating fromother processes, for example as described in the German PatentApplication P 10028582.1.

The output from the reactor is worked up in a manner known per se, forexample by separating off the molecular hydrogen present in the productmixture, separating off constituents other than alkanes and alkenes,preferably by selective absorption of the alkene/alkane mixture in anorganic solvent, and fractionation of the alkene/alkane mixture in a C₃splitter and recirculation of the alkane to the dehydrogenation.

The invention is illustrated by the following examples.

EXAMPLES Example 1

A solution of 11.992 g of SnCl₂.2H₂O and 7.888 g of H₂PtCl₆.6H₂O in 5950ml of ethanol was poured over 1000 g of a granulated ZrO₂.SiO₂ mixedoxide from Norton (screen fraction: 1.6 to 2 mm).

The supernatant ethanol was taken off on a rotary evaporator. The mixedoxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 7.68 g of CsNO₃, 13.54 gof KNO₃ and 98.329 g of La(NO₃)₃.6H₂O in 23 ml of H₂O was then pouredover the catalyst obtained. The supernatant water was taken off on arotary evaporator. The material was subsequently dried at 100° C. for 15hours and calcined at 560° C. for 3 hours.

The catalyst had a BET surface area of 85 m²/g. Mercury porosimetrymeasurements indicated a pore volume of 0.29 ml/g.

Example 2

A solution of 0.6 g of SnCl₂.2H₂O and 0.394 g of H₂PtCl₆.6H₂O in 300 mlof ethanol was poured over 55 g of a granulated ZrO₂.SiO₂ mixed oxidefrom Norton (screen fraction: 1.6 to 2 mm).

The supernatant ethanol was taken off on a rotary evaporator. The mixedoxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.386 g of CsNO₃, 0.680 gof KNO₃ and 4.888 g of Ce(NO₃)₃.6H₂O in 130 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 1 00° C. for 15 hoursand calcined at 560° C. for 3 hours.

The catalyst had a BET surface area of 72.4 m²/g. Mercury porosimetrymeasurements indicated a pore volume of 0.26 ml/g.

Example 3

A solution of 0.684 g of SnCl₂.2H₂O and 0.45 g of H₂PtCl₆.6H₂O in 342 mlof ethanol was poured over 57 g of a granulated ZrO₂ support from Norton(screen fraction: 1.6 to 2 mm).

The supernatant ethanol was taken off on a rotary evaporator. The mixedoxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.44 g of CsNO₃, 0.775 gof KNO₃ and 5.604 g of Ce(NO₃)₃.6H₂O in 148 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

The catalyst had a BET surface area of 40 m²/g. Mercury porosimetrymeasurements indicated a pore volume of 0.25 ml/g.

Example 4

In a 5 l stirred flask, 521.3 g of Zr(OH)₄ were suspended in 2000 ml ofH2O. 73.53 g of SiO₂ Ludox sol (SiO₂ content: 47.6% by weight) wereadded to the suspension. The suspension was stirred at room temperaturefor 4 hours. The product was subsequently spray dried. The temperatureat the top was set to 350° C., the outlet temperature was from 105 to110° C., and the spraying pressure was 2.5 bar. The atomizer diskrotated at a speed of 28 000 rpm. The resulting white powder had a losson ignition of 15.1%.

471.15 g of the white powder were kneaded with 133.30 g of Pural SCF(Al₂O₃) and 30.22 g of concentrated HNO₃ for 2 hours. The paste wasshaped to form 3 mm extrudates by means of a ram extruder (pressingpressure: 75 bar). The extrudates were dried at 200° C. for 4 hours andcalcined at 600° C. for 2 hours. The extrudates were subsequentlycrushed to give particles of a screen fraction from 1.6 to 2 mm.

A solution of 0.712 g of SnCl₂.2H₂O and 0.468 g of H₂PtCl₆.6H₂O in 368ml of ethanol was poured over 60 g of the support produced in this way.

The supernatant ethanol was taken off on a rotary evaporator. The mixedoxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.458 g of CsNO₃, 0.807 gof KNO₃ and 5.838 g of La(NO₃)₃.6H₂O in 157 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

The catalyst had a BET surface area of 98 m²/g. Mercury porosimetrymeasurements indicated a pore volume of 0.35 ml/g.

Example 5

23 g of a granulated Mg(Al)O support from Giulini (screen fraction: 1.6to 2 mm) were calcined at 700° C. for 2 hours. A solution of 0.276 g ofSnCl₂.2H₂O and 0.181 g of H₂PtCl₆.6H₂O in 138 ml of ethanol wassubsequently poured over the support.

The supernatant ethanol was taken off on a rotary evaporator. The mixedoxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.177 g of CsNO₃, 0.313 gof KNO₃ and 2.262 g of La(NO₃)₃.6H₂O in 60 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

The catalyst had a BET surface area of 103 m²/g. Mercury porosimetrymeasurements indicated a pore volume of 0.51 ml/g.

Example 6

A solution of 0.3718 g of SnCl₂.2H₂O and 0.245 g of H₂PtCl₆.6H₂O in 190ml of ethanol was poured over 57 g of a granulated theta-Al₂O₃ supportfrom Condea (screen fraction: 1.6 to 2 mm).

The supernatant ethanol was taken off on a rotary evaporator. The mixedoxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.239 g of CsNO₃, 0.4214g of KNO₃ and 5.604 g of La(NO₃)₃.6H₂O in 80 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

The catalyst had a BET surface area of 119 m²/g. Mercury porosimetrymeasurements indicated a pore volume of 0.66 ml/g.

Example 7

A solution of 0.2758 g of SnCl₂.2H₂O and 0.1814 g of H₂PtCl₆.6H₂O in 138ml of ethanol was poured over 23 g of a granulated theta-Al₂O₃ supportfrom BASF (screen fraction: 1.6 to 2 mm).

The supernatant ethanol was taken off on a rotary evaporator. The mixedoxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.1773 g of CsNO₃, 0.3127g of KNO₃ and 2.26 g of La(NO₃)₃.6H₂O in 60 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

The catalyst had a BET surface area of 34 m²/g. Mercury porosimetrymeasurements indicated a pore volume of 0.23 ml/g.

Example 8

43.25 g of (NH₄)₂CO₃ were dissolved in 1 l of water, mixed with 849 mlof a 25% strength by weight ammonia solution and heated to 75° C. 2333.3g of Mg(NO₃)₂.6H₂O and 337.6 g of Al(NO₃)₃.9H₂O dissolved in 3 l ofwater were added quickly to the solution from a dropping funnel whilestirring. After the mixture had been stirred at 75° C. for 1 hour, theresulting precipitate was filtered off and the filter cake was washedwith water. The solid was subsequently dried at 100° C. for hours andcalcined at 900° C. for 2 hours.

The powder was mixed with 3% by weight of magnesium stearate andprecompacted to form 20×2 mm tablets on an eccentric press.

33 g of the Mg(Al)O support produced in this way were crushed (screenfraction: 1.6 to 2 mm) and calcined at 700° C. for 2 hours. A solutionof 0.398 g of SnCl₂.2H₂O and 0.262 g of H₂PtCl₆.6H₂O in 200 ml ofethanol was subsequently poured over the support.

The supernatant ethanol was taken off on a rotary evaporator. The mixedoxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.256 g of CsNO₃, 0.451 gof KNO₃ and 3.265 g of La(NO₃)₃.6H₂O in 87 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

The catalyst had a BET surface area of 85 m²/g. Mercury porosimetrymeasurements indicated a pore volume of 0.28 ml/g.

Example 9

Catalyst Test:

20 ml of the previously produced catalyst were installed in a tubereactor having an internal diameter of 20 mm. The catalyst was treatedwith hydrogen for 30 minutes at 500° C. The catalyst was then exposed toa mixture of 80% by volume of nitrogen and 20% by volume of air (leanair) at the same temperature. After a flushing phase of 15 minutes usingpure nitrogen, the catalyst was reduced by means of hydrogen for 30minutes. The catalyst was then supplied with 20 standard l/h of propane(99.5% by volume) and H₂O in a molar ratio of propane/water vapor of 1:1at a reaction temperature of 610° C. The pressure was 1.5 bar and thespace velocity (GHSV) was 2000 h⁻¹. The reaction products were analyzedby gas chromatography. The results are summarized in the table.

The Brönsted and Lewis acidities of the catalysts produced in Examples 1to 8 were determined by means of the probe gas pyridine using an HV-FTIRmeasurement cell.

The samples were baked in air at 390° C. for 1 hour, subsequentlyevacuated to 10⁻⁵ mbar, cooled to 80° C. and treated with gaseouspyridine at an equilibrium pressure of 3 mbar. To test the evacuationstability of the pyridine adsorbate, the sample which had been treatedwith gaseous pyridine at 3 mbar was subjected to a vacuum treatment inan oil pump vacuum (about 10⁻² mbar, 3 min) and under high vacuum (about10⁻⁵ mbar, 1 h). Physisorbate material was thus desorbed. The adsorbatespectra were recorded in a high vacuum.

The measured absorbances were expressed as a ratio to the thickness ofthe sample (in integrated extinction units (IEE) per μm of thickness).The single-beam spectrum of the sample which had not been treated withpyridine gas and had been cooled to 80° C. under high vacuum served asbackground for the adsorbate spectra. Matrix bands were thus completelybalanced out.

The band at 1440 cm⁻¹ (corresponds to Lewis-acid centers) andadditionally the control band at 1490 cm⁻¹ (corresponds to Lewis-acidcenters if no Brönsted-acid centers are present) were evaluated.

The results are summarized in the table.

All samples examined displayed no measurable Brönsted acidity. Themeasured Lewis acidity correlates well with the conversion in thedehydrogenation of propane.

TABLE Propane Lewis acidity conversion Example Support (AU) (%) 1ZrO₂/SiO₂ (Norton) 8.97 46.8 2 ZrO₂/SiO₂ (Ce instead of La) 7.82 47.2 3ZrO₂ (Norton) 4.59 30 4 ZrO₂/SiO₂/Al₂O₃ 7.29 42.1 5 Mg(Al)O (Giulini)2.88 12.4 6 theta-Al₂O₃ (Condea) 3.51 30.4 7 theta-Al₂O₃ (BASF) 1.7711.4 8 Mg(Al)O 1.66 12.2 All supports are loaded withPt_(0.3)/Sn_(0.6)/Cs_(0.5)/K_(0.5)/La_(3.0).

Example 10

High-load Operation

2.5 ml of the catalyst produced as described in Example 1 were dilutedwith 77.5 ml of steatite and installed in a tube reactor having aninternal diameter of 20 mm. The catalyst was treated in succession, for30 minutes each at 500° C., firstly with hydrogen, then with lean air(80% by volume of nitrogen and 20% by volume of air) and subsequentlyagain with hydrogen. The catalyst was flushed with nitrogen for 15minutes between each of the treatments. The catalyst was subsequentlysupplied at 600° C. with 20 standard l/h of propane (99.5% by volume)and water vapor in a molar ratio of propane/H₂O of 1:1. The pressure was1.5 bar and the space velocity (GHSV) was 16 000 h⁻. The reactionproducts were analyzed by gas chromatography. After a reaction time ofone hour, 30% of the input propane were converted at a selectivity topropene of 95%. The space-time yield of propene, based on the catalystvolume used, was 8 g of propene/(g of catalyst*h).

Example 11

Operation with Oxygen

20 ml of the catalyst produced as described in Example 1 were installedin a tube reactor having an internal diameter of 20 mm. The catalyst wastreated in succession, for 30 minutes each at 500° C., firstly withhydrogen, then with lean air (80% by volume of nitrogen and 20% byvolume of air) and subsequently again with hydrogen. The catalyst wasflushed with nitrogen for 15 minutes between each of the treatments. Thecatalyst was subsequently supplied at 610° C. with 20 standard l/h ofpropane (99.5% by volume) and water vapor in a molar ratio ofpropane/H₂O of 1:1. In addition, oxygen was introduced in a molar ratioof propane/O₂ of 20:1. The pressure was 1.5 bar and the space velocity(GHSV) was 2100 h⁻¹. The reaction products were analyzed by gaschromatography. After a reaction time of one hour, 50% of the inputpropane were converted at a selectivity to propene of 90%. After areaction time of 16 hours, the conversion was 44% and the selectivitywas 90%.

Example 12

Operation with Oxygen at Low Conversion

20 ml of the catalyst produced as described in Example 1 were installedin a tube reactor having an internal diameter of 20 mm. The catalyst wastreated in succession, for 30 minutes each at 500° C., firstly withhydrogen, then with lean air (80% by volume of nitrogen and 20% byvolume of air) and subsequently again with hydrogen. The catalyst wasflushed with nitrogen for 15 minutes between each of the treatments. Thecatalyst was subsequently supplied at 500° C. with 20 standard l/h ofpropane (99.5% by volume) and water vapor in a molar ratio ofpropane/H₂O of 1:1. In addition, oxygen was introduced in a molar ratioof propane/O₂ of 20:1. The pressure was 1.5 bar and the space velocity(GHSV) was 2100 h⁻¹. The reaction products were analyzed by gaschromatography. After a reaction time of one hour, 16% of the inputpropane were converted at a selectivity to propene of 99%. After areaction time of 100 hours, the conversion was 14% and the selectivitywas 94%. After increasing the temperature to 510° C., the propaneconversion after 300 hours was 15% and the selectivity was 94%. Afterincreasing the temperature further to 530° C., 15% of the input propanewere being converted at a selectivity to propene of 94% after 800 hours.After 1700 hours, the supply of propane and of water was stopped andlean air (80% by volume of nitrogen and 20% by volume of air) was passedover the catalyst at 400° C. Pure air was then passed over the catalystfor 30 minutes. After the reactor had been flushed with nitrogen for 15minutes, hydrogen was passed over the catalyst for 30 minutes. Afterpropane, water vapor and oxygen were again supplied as feed, a propaneconversion of 15% at a selectivity of 92% could be achieved at 505° C.After a total of 2300 hours, the propane conversion at 540° C. was 15%and the selectivity to propene was 94%.

Example 13

Operation with Oxygen at Low Conversion Using Additional N₂ Dilution

After 2300 hours, the catalyst from Example 1 was again (after stoppingpropane and water vapor supply) treated with lean air (80% by volume ofnitrogen and 20% by volume of air) at 400° C. Subsequently, pure air waspassed over the catalyst for 30 minutes. After the reactor had beenflushed with nitrogen for 15 minutes, hydrogen was passed over thecatalyst for 30 minutes. Subsequently, propane, nitrogen, oxygen andwater vapor in a ratio of 5.8/7.8/0.4/5.8 were passed over the catalystat 505° C. The reaction pressure was 1.5 bar and the space velocity(GHSV) was 1300 h³¹ ¹. The propane conversion was 20% at a selectivityof 92%. After 500 hours, 20% of the propane were converted at aselectivity to propene of 92%.

Example 14

Operation with Oxygen at Low Conversion Using Additional N₂ Dilution andAdditional Introduction of H₂

Using the experimental procedure of Example 12, hydrogen wasadditionally mixed into the feed after a total running time of 2500hours. The feed then had the following composition:C3/N₂/O₂/H₂/H₂O=5.8/7.8/0.4/0.8/5.8. The reaction pressure was 1.5 barand the space velocity (GHSV) was 1300 h⁻¹. The reaction temperature wasset to 575° C. The propane conversion was 20% at a propene selectivityof 92%. The oxygen introduced was reacted completely. 60% of the oxygenintroduced reacted with propane or propene to form carbon dioxide andcarbon monoxide, 40% of the oxygen introduced reacted with hydrogenwhich had been introduced or formed in the dehydrogenation to givewater.

Example 15

Regeneration of the Catalyst

1000 ml of the catalyst produced as described in Example 6 were dilutedwith 500 ml of steatite and installed in a tube reactor having aninternal diameter of 40 mm. The catalyst was treated in succession, for30 minutes each at 500° C., firstly with hydrogen, then with lean air(80% by volume of nitrogen and 20% by volume of air) and subsequentlyagain with hydrogen. The catalyst was flushed with nitrogen for 15minutes between each of the treatments. The catalyst was subsequentlysupplied at 610° C. with 250 standard l/h of propane (99.5% by volume)and water vapor in a molar ratio of propane/H₂O of 1:1. The pressure was1.5 bar and the space velocity (GHSV) was 500 h⁻¹. The reaction productswere analyzed by gas chromatography. After a reaction time of one hour,55% of the input propane were converted at a selectivity to propene of90%. After a reaction time of 12 hours, the conversion was 53% and theselectivity was 93%. The supply of propane and water was stopped andlean air (92% by volume of nitrogen and 8% by volume of air) was passedover the catalyst at 400° C. The air content was subsequently increasedtwice (firstly to 83% by volume of nitrogen and 17% by volume of air,then to 64% by volume of nitrogen and 36% by volume of air). Pure airwas then passed over the catalyst until the CO₂ content of the outputgas was less than 0.04% by volume. After the reactor had been flushedwith nitrogen for 15 minutes, hydrogen was passed over the catalyst for30 minutes. After propane, water vapor and oxygen were again supplied asfeed, a propane conversion of 55% at a selectivity of 92% could beachieved at 610° C. After the catalyst had been regenerated 10 times inthe above-described manner, a conversion of 54% at a propene selectivityof 93% could be achieved at 610° C. After the catalyst had beenregenerated 30 times, a conversion of 54% at a propene selectivity of93% could be achieved at 610° C.

1. A process for the heterogeneously catalyzed dehydrogenation in one ormore reaction zones of one or more dehydrogenatable C₂-C₃₀-hydrocarbonsin a reaction gas mixture comprising them, with at least part of theheat of dehydrogenation required being generated directly in thereaction gas mixture in at least one reaction zone by combustion ofhydrogen, the hydrocarbon or hydrocarbons and/or carbon in the presenceof an oxygen-containing gas, wherein the reaction gas mixture comprisingthe dehydrogenatable hydrocarbon or hydrocarbons is brought into contactwith a Lewis-acid dehydrogenation catalyst which has essentially noBronsted acidity, wherein the dehydrogenation catalyst has a Lewisacidity of greater than 3 acidity units (AU), determinable from IRabsorption spectra of pyridine adsorbed on the catalyst.
 2. A process asclaimed in claim 1, wherein the dehydrogenation catalyst comprises ametal oxide selected from the group consisting of zirconium dioxide,aluminum oxide, silicon dioxide, titanium dioxide, magnesium oxide,lanthanum oxide and cerium oxide.
 3. A process as claimed in claim 2,wherein the dehydrogenation catalyst comprises zirconium dioxide and/orsilicon dioxide.
 4. A process as claimed in any of claim 1, wherein thedehydrogenation catalyst comprises at least one element of transitiongroup VIII, at least one element of main group I or II, at least oneelement of main group III or IV and at least one element of transitiongroup III including the lanthanides and actinides.
 5. A process asclaimed in any of claim 1, wherein the dehydrogenation catalystcomprises platinum and/or palladium.
 6. A process as claimed in any ofclaim 1, wherein the dehydrogenation catalyst comprises cesium and/orpotassium.
 7. A process as claimed in any of claim 1, wherein thedehydrogenation catalyst comprises lanthanum and/or cerium.
 8. A processas claimed in any of claim 1, wherein the dehydrogenation catalystcomprises tin.
 9. A process as claimed in any of claim 1, wherein thedehydrogenation catalyst has a bimodal pore radius distribution in whichfrom 70% to 100% of the pores have a pore diameter less than 20 nm or inthe range from 40 to 5000 nm.
 10. A process as claimed in any of claim1, wherein the reaction gas mixture comprises water vapor.
 11. A processas claimed in any of claim 1, wherein hydrogen is added to the reactiongas mixture.
 12. A process as claimed in claim 11, wherein at least onereaction zone contains a catalyst which selectively catalyzes thecombustion reaction of hydrogen and oxygen in the presence ofhydrocarbons.
 13. A process as claimed in any of claim 1, wherein thecatalyst which catalyzes the combustion of hydrogen comprises oxides orphosphates selected from the group consisting of the oxides andphosphates of germanium, tin, lead, arsenic, antimony and bismuth.
 14. Aprocess as claimed in claim 1, wherein the catalyst which whichselectively catalyzes the combustion of hydrogen comprises a noble metalof transition group VIII or I.
 15. A process as claimed in claim 1,wherein the dehydrogenation is carried out in a tray reactor.
 16. Aprocess as claimed in claim 1, wherein the dehydrogenation catalyst hasa Lewis acidity of greater than 6 acidity units (AU), determinable fromIR absorption spectra of pyridine adsorbed on the catalyst.
 17. Aprocess as claimed in claim 16, wherein the dehydrogenation catalystcomprises a metal oxide selected form the group consisting of zirconiumdioxide, aluminium oxide, silicon dioxide, titanium dioxide, magnesiumoxide, lanthanum oxide and cerium oxide.
 18. A process as claimed inclaim 17, wherein the dehydrogenation catalyst comprises zirconiumdioxide and silicon dioxide.